With long on-stream period



March 31, 1964 J, w. scoTT, JR

HYDROCARBON HYDROCRACKING PROCESS WITH LONG ON-STREAM PERIOD Filed Deo. 28, 1961 United States Patent O HYDRCARBN HYDRGCRACKING PROCESS lohn W. Scott, Jr., Ross, Calif., assigner to California Research Corporation, San Francisco, Calif., a corporation of Delaware Filed Dee. 28, 1961, Ser. No. 165,059

1 Claim. (Cl. 208-111) This invention relates to a catalytic process for upgrading a vvariety of petroleum fractions into premium gasolines of exceptionally high octane rating and with good boiling point characteristics. More particularly, the process -is one whereby naphthas, non-residual cycle oils from thermal or catalytic cracking units, and Minas and other para'nic-type residual stocks can be converted in good yield to premium gasolines of exceptionally high octane rating and which are substantially free of undes1rable higher boiling components.

This application is a continuation-impart of application Serial No. 837,991, filed September 3, 1959, which is a continuation-in-part of application Serial No. 612,802, filed September 28, 1956, which, in turn, is a continuation-in-part of application Serial No. 497,922, filed March 30, 1955, and now all abandoned.

The manufacturer of automotive engines of high compression ratio places the petroleum rener under obligation to provide at reasonable cost a gasoline which is sufciently high in octane rating and other required characteristics as to give optimum engine performance. With the advance of compression ratios into the range of 8-8.5 to 1, motor requirements have been satisfactorily met by gasoline having leaded octane ratings of from about 90 to 95. As the compression ratio has advanced beyond 8.5 to l, the petroleum industry has been able to satisfy engine requirements by resort to Widespread introduction of expensive catalytic cracking and catalytic reforming units capable of providing gasolines having leaded octane ratings as high las 97 or even 98. However, even with these facilities, it has not been possible to obtain large yields of the desired gasolines from readily available starting materials. Instead, each unit or process facility capable of converting only a selected feedstock, and itis not possible to obtain products of the desired quality with any one of a wide variety of feed materials. n

ln addition to the feedstock limitation, the rener/ 1ndustry is now faced with the imminent introduction of automotive engines having compression ratios of 10 to 1 and above which will require fuels having leaded octane ratings of 100 or more. Stocks of this character can currently be produced only by selecting limited fractions from various catalytic reforming or isomerization units, and by resort to expensive `alkylation procedures leading to the production of specialized fuels of the type employed in aircraft engines. As a result, the feel-ing is now general in the refinery industry that such gasolines cannot be produced in large volumes except at costs well above those currently associated with premium gasolines.

It is, therefore, a general object of this invention to provide a process whereby widely available petroleum fractions of relatively low octane rating can be converted in good yields to gasolines having leaded octane ratings as high as 100 or more, if desired, and which are relatively free of components boiling above about 350 to 400 F. A more particular object is to obtain gasolines of this character from a variety of fee'dstoclcs available in ample volume to permit a large scale gasoline production. The nature of still other, and more particular objects of the invention will be apparent from a consideration of the descriptive portion to follow.

The present invention is based on 'the discovery that 1t is possible to convert relatively low nitrogen content Iraphthas, non-residual cycle oils, and residual stocks of mineral oil origin to gasolines `of unusually high octane rating by passing the feedstock, along with at le-ast 1500 s.c.f. of hydrogen per barrel of said feed, continuously for an on-stream period of at least several thousand hours, through an isomeriZation-cracking zone provided With an active catalyst made up of one or more hydrogenatingdehydrogenating components, including nickel sulde, intimately associated with a support having active cracking characteristics, said Zone being maintained at pressures of at least 600 p.s.i.g. and `at temperatures below 700 F. and preferably between about 400 and 650 F. during at least the first half of the on-stream period, starting the reaction at a temperature below 600 F., and by raising the temperature of said Zone to successively higher values during said first half of the on-stream period, and preferably during the entire ori-stream period, as necessary to maintain substantially constant conversion of above 50 volume percent. The space rate utilized may be varied within relatively wide limits, eg., from 0.1 to 15. However, a preferred range is from about 0.2` to 5. [if desired, temperatures above 700 F., up to about 875 F., can be employed during any remaining portion of the onstrearn period in order to maintain adequate conversion levels, though this entails increased catalyst fouling rates, increased gas make, decreased liquid yields, and lower ring yields. It has been found that there is a catalyst temperature level at which the catalyst fouling rate suddenly 'and dramatically become-s exponential instead of linear, with feeds of any nitrogen content. This point is reached sooner with feeds `containing more than 200 p.p.m. nitrogen than with feeds containing less nitrogen. However, with feeds containing less than 200 ppm. nitrogen, the critical temperature at which the fouling rate suddenly becomes exponential still is reached too soon for adequate run lengths unless the starting temperature is kept below 600 F. and unless at least the first half of the ori-stream period is conducted at average catalyst temperatures below 700 F. Substantially constant conversion is necessary in order to provide reasonably practical operation including adequate throughput. Substantially constant conversion could be achieved by reducing space velocity instead of increasing temperature as the catalyst fouls; however, this would cause a generally intolerable or undesirable reduction in throughput. The reaction occurring during operation of the process of the present invention results in Ia substantial net consumption of hydrogen as the feed (including any recycled stock) is converted to the desired high octane gasoline product fractions vvhich are recovered from the etliuent from the isomeriZation-cracking zone. Said product fractions boil essentially below the initial boiling point of the feed and can therefore `be referred to as synthetic products.

ln the preferred practice of this invention, the eluent from the foregoing isomerization-cracking zone is freed of lighter, normally gaseous components and is then fractionated to recover a light fraction comprising high octane gasoline, an intermediate gasoline fraction which is also of good `octane value, and a bottoms fraction (normally boiling above Iabout 300 to 400 F.) which can be recycled to -the isomerization-cracking Zone for further conversion to la synthetic gasoline pnoduct of desired end point. The intermediate fraction can Ibe upgraded in octane value still further by passing the same, along with added hydrogen, through a catalytic reforming zone under conventional reforming conditions (representant-'e reforming temperatures and pressures being from yabout 850 to 1000 F. and from about 200 to 900 p.s.i.g.), and this manner of processing is most preferred, for, quite surprisingly, it has been found that the effluent from the reforming zone, after being lfreed of normally gaseous components, is a gasoline of higher octane rating than could be obtained by a practice of either the reforming step yalone or by using only the isomerizat-ion-cracking step.

As stated above, the present invention is adapted to the 4conversion of mineral oil naphthas and cycle oils, including those derived from petroleum, shale or gilsonite sources, for example, `as well as Minas and other paraiiinic-type -residual stocks. Representative stocks are petroleum naphthas of straight run, catalytic `or thermal cracked origin and boiling in a range of `from about 175 to 500 F. Other representative :feedstocks are nonresidual petroleum cycle oils, preferably boiling below about 750 F., yas obtained from the effluent from a thermal or `a catalytic cracking unit or from la coking unit. O-f `such stocks, those of lighter character boiling below about 650 F. and preferably between about 360 and about 600 F. have been found to yield gasolines of an unusually high octane rating when processed in accord ance with the present invention. The references herein to boiling points or ranges .are those measured in accordance with either of the ASTM distillation procedures D-86 or D-15S, depending on boiling range. It is further -to be understood that, in respect to the designations of boiling ranges and feed and product distillation cut points, a by volume tolerance is to be permitted in order to more closely approximate the practical limitations of refinery ydistillation equipment and practices. Thus, the ydesignation of a feed boiling range between 360 and 600 F. resolves itself to these feeds wherein at least the 10% and 90% distillation points fall 4within the stated range.

While it is evident from the foregoing paragraph that the benefits of the present invention can be obtained with diverse feedstocks, it is necessary that the latter be relatively low in total nitrogen if poisoning of the isomerization-cracking catalyst is to be avoided. rlhus, the stock employed should be one containing less than 200 p.p.m. of nitrogen `and preferably less than 50 ppm. of this component, it being recognized that this specification is somewhat flexible, since higher nitrogen contents (those approaching 200 p.p.m.) can be tolerated with heavier, higher boiling stocks than with those boiling essentially in the gasoline range. ln the case of stocks which are not already sufficiently low in nitrogen, acceptable levels can be reached by pretreating the feed material with hydrogen at elevated temperatures and pressures in the presence of a suitable catalyst which is relatively free of cracking characteristics. A particularly effective catalyst for this purpose is one wherein a coprecipitated molybdena-alumina material is combined with cobalt oxide, the lfina-l catalyst having a metals content of about 2% cobalt and 7% molybdenum. Representative processing conditions 'for removing nitrogen with this catalyst are an LHSV of 1, 700-800 F., 20G-1500 p.s.i.g. and 1000- 15,000 s.c.f. of hydrogen per barrel of feed-stock. The effluent vfrom this pretreating, or hydro-raining, step can either be fed directly to the isomerization-cracking zone, as discussed more fully below, or it can be lirist subjected to a preliminary fractionation to remove the small amounts of low octane gasoline components which are formed `during more vigorous denitriication treatments.

The catalyst employed in the isomerization-cracking Zone is one wherein a material having hydrogenating-dehydrogenating activity is deposited or otherwise `dispos-ed on an 4active cracking catalyst support. The cracking component may comprise any one or more of such acidic, synthetically prepared materials as silica-alumina, silicamagnesia, silica-alumina-zirconia composites, EP3 on alumina, BFS on silica-alumina, various `acid-treated clays, and similar materials. The hydrogenating-dehydrogenting components of the catalyst can `be selected from the various group Vi and group Vlll metals, as well as from the oxides and suldes thereof, representative materials being the oxides and sulfides of molybdenum, tungsten, chromium, and the like, as well as metals such as nickel or cobalt and various oxides and sullides thereof. Also suitable are various groups IB or l'lB metals7 such as copper or cadmium and their oxides and sullides. If desired, more than one hydrogenating-dehydrogenating component can be present, :and good results have been obtained with catalysts containing composites of two or more of the oxides or sullides of molybdenum, cobalt, nickel, chromium and Zinc. Depending on the activity thereof, the `amount of the hydrogenating-dehydrogenating component present can be varied within relatively wide limits of from about 0.1 -to 30%, based on the weight of the entire catalyst. Within these limits, the yamount of said component present should be suflicient to provide a reasonable catalyst onstream period at required conversion levels, but insufficient to effect Substantial saturation of any except highly substituted and polynuclear aromatic -rings under the reaction conditions employed `in the isomerization-cracking zone.

yParticularly good results from `the Istandpoint of high per-pass conversion, even at relatively low operating temperatures, coupled with good selectivity and the ability to withstand repeated regeneration with relatively minor decrease in activity, `are obtained with catalysts composed of from 1 to 25% nickel sulfide deposited on synthetically prepared silica-alumina composites containing from about to 95% silica.

The catalyst employed in the reforming Zone can be any one or more of the various materials which are now available for effecting reforming operations. One or the other of two catalysts are commonly employed in such reactions, and either one or both of the same can be employed in the practice of the present invention. One of these catalysts comprises about 8 to 12% molybdenum oxide disposed on an alumina support, While the other usually contains from about 0.1 to 1.0% by weight of metallic platinum dispersed on an alumina support, along With small amounts of halogens. The latter catalyst is commonly referred to in the art as Platforming catalyst.

Having selected or prepared the (low nitrogen) feedstock to be employed as well as the desired catalyst components, the process of the present invention can be carried out in a continuous fashion by preheating the feed and added hydrogen to the extent necessary to bring the reaction in the isomerization-cracking zone on stream at the desired temperature, which must be in the range of from 400 to 600 F. Thereafter, while maintaining pressure and feed throughput rates at the desired levels, the temperature in the said zone is controlled so as to maintain (average) reaction temperatures therein at a level below 700 P. during at least the first half of the onstream reaction period, and preferably during the entire duration of said period. The temperature is continuously or periodically raised from the 450 F. to 600 F. starting temperature as the reaction progresses, at a rate which will maintain, during at least the first half of the onstream period and preferably during the entire on-stream period, substantially constant conversion of the feed to products boiling below the initial boiling point of the feed. While the duration of the on-stream period will vary from stock to stock, under favorable operation conditions the isomerization-cracking catalyst remains active for several thousand hours, with activity being then restored by a conventional regeneration involving the burning of contaminants with an oxygen-containing gas, if desired.

rl`he effluent from the isomerization-cracking zone is freed of hydrogen and other normally gaseous components and is then passed to a fractionation zone for separation of the desired synthetic gasoline product. Preferably the separation is so effected as to supply a light gasoline fraction (EP. about 16S-200 E), an intermediate fraction having an end point of about 300400 F. which preferably is fed to the reforming Zone, and a bottoms fraction which is normally returned as4 recycle to the isomerizationcracking zone. In some cases, as with lighter feedstocks, or when the Conversion in the isomerization-cracking zone is unusually high, the eiliuent from this zone will contain little if any material boiling above the so-called intermediate fraction. In this case, recycle from the fractionation zone to the isomerization-cracking Zone is omitted, with the normally liquid effluent from the latter zone being divided into a light product gasoline fraction having an end point of from 165 to 200 F., and a heavy fraction which is then preferably passed to the reforming zone if maximum upgrading of octane rating is to be obtained.

In the reaction occurring in the isomerization-cracking zone, there is a net consumption of hydrogen normally amounting to from about 1000 to 2000 s.c.f. of hydrogen per barrel of feed converted in said zone. In this connection, conversion is arbitrarily expressed in terms of the amount of feed (including recycled eiliuent bottoms) converted to product boiling below the initial boiling point of said feedstocks.

The intermediate fraction referred to above, normally boiling between about 175 l10 F. and 350150o F., may be recovered as product, though it is preferably passed through a furnace or other heat exchange unit where it is heated to a temperature of from about 800 to 1000 F. (preferably 850 to 950 F.), with the resulting heated stream, along with from about 2000 to 10,000 s.c.f. of hydrogen per barrel of said fraction, being passed at a pressure of from 200 to 900 p.s.i.g. (preferably 350 to 600 p.s.i.g.) through the catalyst in the reforming zone at an LHSV of from about 0.1 to 5, and preferably of from l to 3. The effluent from the reforming zone is then freed of hydrogen and other normally gaseous components, including C3 and preferably C4 hydrocarbons as well. The hydrogen, of which there is normally a net production of from about 600 to 1000 s.c.f. per barrel of feed to the reforming zone, is recycled to one or the other of the catalyst zones, while the balance of the product represents gasoline having an octane rating of 100 F-1-{3 ml. or less of TEL, and which in many cases is 100+F-1 clear. This gasoline fraction has boiling point characteristics similar to those of the feed to the reforming zone and is thus substantially free of higher boiling actions. However, any such higher boiling components in the gasoline can be converted to high octane materials of lower boiling point by recycling said components to the isomerization-cracking zone.

It will be seen from the above that in carrying out the present invention there is normally a net deficiency of hydrogen. This deficiency can be met in any desired way, either by supplying hydrogen in the required amounts from some external source, or by employing as a supplementary feed to the reforming Zone a straight-run gasoline which is rich in naphthenes and thus will be productive of the desired additional amounts of hydrogen.

The yields of unfinished gasoline obtained by a practice of the present invention are extremely high and normally range from about 80 to 110 volume percent in terms of the volume of (liquid) fresh feed supplied to the isomerization-cracking Zone. From this it is obvious that the process is an extremely eicient one entailing little degradation of the feed to less valuable, normally gaseous components.

The manner in which the present invention is practiced can be illustrated by reference to the gure of the appended drawing which is a simplied flow scheme of a relinery unit suitable for use in practicing the invention.

An exemplary arrangement of process units and iiow paths suitable for carrying out the present invention is shown in the drawing. An exemplary operation of the invention will now be discussed in connection with the drawing.

A hydrofined naphtha feedstock is selected, originally representing a 50/50 blend as obtained from a thermal cracker and a catalytic cracker, both operating on crude i? gas oils of California origin, said feed having the following specifications:

The above feed, along with 6000 s.c.f. of hydrogen per barrel of feed as supplied through line 111, is heated to a desired temperature between 450 and 600 F. by passage through heat exchanger 12, the resulting vaporous mixture then being passed at said temperature, at a pressure of 1200 p.s.i.g., and a liquid hourly space velocity (LHSV) of 1 through an isomerization-cracking zone 13 provided with a catalyst made up of molybdenum oxide (7.3% by weight as Mo) and cobalt oxide (2.3% by weight as Co), deposited on a synthetic silica-alumina gel cracking support (TCC beads containing approximately 87% silica, 13% alumina). Alternatively, instead of cobalt oxide and molybdenum oxide, nickel sulde (e.g., 6% by weight as nickel) may be used with said support. As the run progresses, the average catalyst temperature is continuously or periodically raised as necessary to maintain substantially constant conversion of the feed to products boiling below the initial boiling point of the feed. The effluent stream from zone 13 is passed through line i4 and cooler 114 into gas-liquid separator 15 from which a gas stream made up essentially of hydrogen is taken through line 16 and recycled to zone 13 via lines 17 and 11. A liquid phase is withdrawn from separator 15 through line 1S and passed into a second gas-liquid separator 19, operated at essentially atmospheric pressure, from which a gaseous stream made up of C4 and lower boiling components is taken overhead and out of the system through line 20. The liquid stream from separator 19 is then fed through line 21 to a fractionation zone 22 where the stream is distilled into a gasoline fraction having an end point of about 165 F. and taken through line 23, an intermediate fraction boiling from about 165 to 300 F. and taken through line 26, and a bottoms fraction boiling above 300 F. and recycled via line 24 to line 10 for passage through isomerizationcracking zone 13.

The gasoline fraction in line 213 has octane rating of 86-88 F-l clear and up to 1.00 F-l-l-3 ml. TEL. rThe catalyst is preferably regenerated in the conventional fashion by burning with air as the per-pass conversion falls much below about 30%.

The intermediate fraction from zone 22 in line 26 (which fraction has an octane rating of 7 8, F-.l clear and of 92.2 F-l with 3 ml. TEL), along with 6000 s.c.f. of hydrogen per barrel of the feed fraction as supplied through line 27, is then heated to a temperature of 850 F. by heat exchanger ZS and passed at a pressure of 500 p.s.i.g. into reforming zone 29 at an LHSV of 2A The catalyst in Zone 29 may comprise 0.3 platinum disposed on an alumina support. The effluent from zone 29 is then passed through line 30 and cooler 30 to a gas-liquid separator 31, operated at approximately 500 p.s.i.g. From this separator a gas stream consisting essentially of hydrogen is taken overhead through line 17 for return to the isomerization-cracking and reforming zones through lines 11 and 27, said hydrogen lines also being supplied with fresh hydrogen from line 11 in the amount of about 1200 s.c.f. per barrel of fresh feed to line to make up for the hydrogen deficiency in the system, the amount of hydrogen consumed in zone 13 being this amount greater than that produced in the reforming zone 29. When it is not desired to supply hydrogen to the system in this fashion, much the same result can be obtained by supplying an auxiliary feed through line 3?. to the reforming zone 29, which feed is a material such as straight-run gasoline which is high in naphthenes and is thus productive of relatively large amounts of hydrogen during the catalytic reforming step, the amount of auxiliary feed supplied in this fashion normally being so adjusted that the system is essentially in hydrogen balance.

The liquid fraction from separator 31 is taken through line 33 to gas-liquid separator 34 operated at essentially atmospheric pressure, there being recovered from this separator a normally gaseous, predominantly (2l-C4 fraction which is removed through lines 35 and 20, and a gasoline fraction taken through line ,36. This gasoline fraction, which has essentially the same boiling range as the feed in line 26, has octane ratings of 93.2 F-l clear and greater than 100 F-l-l-3 ml. TEL.

The combined C5-300 F. gasoline product from lines 23 and 36, representing a yield of 82 volume percent, based on fresh feed to line 10, and as obtained over an S-hour onstream period, has an octane rating of at least 100 F-l-l-3 ml. TEL.

Example In this operation there was employed as feed stock a light catalytic cycle oil of California origin of the type having the following specifications:

Gravity, API 25.5 Aniline point, F 75 Aromatics, vol. percent 58 Basic nitrogen, p.p.m 222 Boiling range, ASTM D158, F.:

Start 410 459 30% 472 50% 480 70% 492 90% 511 End point 549 In order to reduce the nitrogen content of this feed to a satisfactory level, the feed was subjected to a hydro- -fining operation wherein the stock was passed at an LHSV of 0.8, a temperature of 730 7 F. and a pressure of 720 p.s.i.g., along with 6000 s.c.f. of hydrogen per barrel of feed, through a catalyst containing 10.5 weight percent molybdenum oxide and 3.7 weight percent cobalt oxide deposited on an alumina support. The resulting material was thereafter treated so as to remove hydrogen, hydrogen sulfide, ammonia and other gases, leaving a hydrolined stock having the following specifications:

Gravity, API 29.6 Aniline point, F 84 Aromatics, vol. percent 54 Basic nitrogen, p.p.m 0.2 Boiling range, ASTM, D-158, F.:

Start 398 10% 440 30% 457 50% 472 70% 484 90% 507 End point 566 The above hydrofined stock was then fed, along with 6500 s.c.f. recycle gas (containing in excess of 90% hydrogen) per barrel of feed, through the isomerizationcracking zone at a starting temperature of about 550 F. and the temperature was continuously raised during a continuous on-stream run of 1300 hours, as necessary to maintain substantially constant conversion of 60 volume percent, i.e., 60 volume percent of the feed was n tra Total Elapsed Average Catalyst Ori-stream Time Temperature (hrs.) F.)

The temperature at the end of the run was 625 F., and the run was conducted at a pressure of 1200 p.s.i.g. and a space rate of 0.8 LHSV. The catalyst in said zone comprised nickel-sulfide (6% by weight nickel) deposited on synthetic silica-alumina.

This catalyst was prepared by impregnating small (143 dia.) beads of a synthetically prepared silica alumina (10%) cracking support with a solution of nickel nitrate having a concentration such as to provide the finished catalyst with the equivalent of 6 Weight percent nickel. The thus impregnated support was then dried at 250 F. and calcined at 1000 F. for l0 hours. This was followed by an added heating step wherein the catalyst was held in a stream of air at temperatures ranging from 1000 to 1400 F. for several hours. The resulting nickel oxide-containing catalyst was then transferred to the reactor unit for reduction and suliiding prior to being placed on stream. Specifically, the catalyst was heated in nitrogen to 570 F., following which hydrogen at this temperature, first at ambient pressure and then at 1200 p.s.i.g., was passed over the catalyst to reduce the nickel oxide to metallic nickel. The hydrogen stream was then admixed with hexanes and isopropyl mercaptan, the concentration of the mercaptan being such as to provide the equivalent of 2 mol percent H28 in hydrogen. This was continued until the unit had been provided with a two-fold excess of HES over that theoretically required to convert the nickel to the sulde form.

In this operation, it was found that the feed was converted to a C25-400 F. product in a yield of 103.5 volume percent, based on feed. A typical product inspection of the said C5-400 F. product is as follows:

Gravity, API 54.2 Aniline point, F 97.5 Parans, vol. percent 38 Naphthenes, vol. percent 38 Aromatics, vol. percent 24 Octane rating, F-l-l-S ml. TEL 88 The C5-180" F. portion of the above C5-400 F. product had an F-l rating of 99 (with 3 ml. TEL). The remainder of the product, i.e., the ISO-400 F. portion is found to have an F-l octane rating (plus 3 ml. TEL) of about 101.9 after being passed through a catalytic reforming unit under conventional reforming conditions, as noted above in the example, said reformed product being produced in a yield of about 90 volume percent, based on feed to the reformer. Blending this reformed product with the lighter C5-l80 F. fraction noted above gives a final (l5-400 F. gasoline product having an F-l octane rating (plus 3 ml. TEL) of 101.1. All portions of the etiiuent from the isomerization-cracking zone boiling above about 400 F. were recycled to said zone so as to convert all portions of the feed to a synthetic product, i.e., one having an end point below the initial boiling point of the feed.

It will be noted that the respective feed streams to the isomerization-cracking Zone are converted in excellent yield to products boiling essentially below the initial boiling point of the said feed streams, and this in general constitutes the preferred manner of practicing the present invention. Such conversion is particularly advantageous in the case of aromatic compounds boiling above about 350-375" F. (e.g., such as found in the bottoms fraction obtained from the reforming zone eluent) since such compounds are not clean-burning and tend to build up engine deposits which increase the octane requirements of the engine. The present conversion method not only converts these compounds (as well as nonaromatic compounds) to lower boiling, more desirable compounds from the burning standpoint, but also effects such conversion while actually increasing the octane number of the product in appreciable measure.

Among the reactions of aromatics which take place in the isomeriza-tion-cracking zone, the principal reaction is evidently fone 'of disproportionation followed by hydrogenation and cracking of the heavier disproportionation products. Thus, trimethyl benzenes in the feed to the isomerization-cracking zone are isomerized and disproportionated to produce xylenes and lighter aromatics on the one hand, and tetramethyl and Ihigher benzens on the other. These heavier products appear to be selectively hydrogenate'd in their ring portion, with the resulting molecule then cracking into relatively light isoparains of high octane rating. In the case of polynuclear aromatic components in the feed to the isomerization-cracking zone, one of the rings is believed to be hydrogenated and fractured, with the resulting substituted aromatic product being thereafter cnacked (or rst disproportionated and then cracked) into the desired light isoparains. While a certain amount of n-parains are also produced during these reactions, the isoparaflin/n-paraflin ratio in the eluent from the isomerization-cnacking zone is well on the iso side of thermodynamic equilibrium due to the aforesaid selective hydrodecomposition of the more highly substituted aromatics and polynuclear components.

Isoparains in the feed are cracked to fractions of lower boiling point, while paraflins in the feed of essentially straight-chain conflgunation are rst isomerized and then immediately cracked. Here again, however, the isoparain/n-parain ratio of the cracked products is far on the iso iside of thermodynamic equilibrium, a fact which is believed to be largely responsible for the unusually high octane number of the product obtained by a practice of this invention. On the other hand, parainic components not so cracked in a given pass through the isomerizationcracking zone |appear in the effluent substantially unchanged as determined by their octane number and boiling point characteristics. While such components are eventually converted on being recycled, this shofws that isomerization in the usual sense of the word is not responsible for the increase in `octane number obtained even in the case of passing nonaromatics through the isomerization-cracking zone.

I claim:

In a process for hydrocracking la hydrocarbon feed stock to produce gasoline, which comprises passing said stock along with at least 1500 s.c.f. of hydrogen per barrel thereof, through a catalyst comprising at least one Ihydrogenating-dehydrogenating component intimately associated rwith an active cracking catalyst support in a hydrocracking zone for a given on-stream period, said zone being maintained at a pressure of atleast 600 p.s.i.g., and recovering said gasoline from the etiluent from said zone, the improvement which comprises: (a) selecting as said feed stock a stock containing from l0 to 50 parts per million nitrogen; (b) initiating said hydrocracking `at 'the beginning of said ion-stream period at a temperature between 400 and 600 F.; (c) using as said catalyst a catalyst having a hydrogenating-dehydrogenating component comprising nickel sulfide; (d) carrying out said hydrocracking at a substantially constant conversion of said `feed of above volume percent to products boiling below the initial boiling point `of said feed; (e) continuing said on-stream period for at least several thousand hours; (f) continuing said hydrocracking at an Iaverage temperature between about 400 and 700 F. during at least the first half of said on-stream period; (g) continuing said hydrocracking at an average temperature between about 700 and 875 F. during any remaining portion of said period; (h) raising said average temperature as necessary during lat least said iirst half Aof said period to maintain said substantially constant conversion; recovering from the ieiluent from said hydrocracking zone said gasoline having a boiling range essentially below the initial boiling point of said feed and having a high octane rating.

References Cited in the tile of this patent UNITED STATES PATENTS 2,651,597 Corner et al. Sept. 8, 1953 2,769,769 Tyson Nov. 6, v1956 2,911,356 Hanson Nov. 3, 1959 

